Process and apparatus for contacting hydrocarbons with catalyst

ABSTRACT

A reactor apparatus and process for contacting hydrocarbons with catalyst. The reactor apparatus comprises a plurality of tubular reactors each having a first end into which a catalyst is fed and a second end through which the catalyst and product exit the tubular reactor. A catalyst retention zone is provided to contain catalyst and feed catalyst to the tubular reactors. A separation zone is provided to separate the catalyst from products of a reaction conducted in the apparatus. A transport conduit having a first end in fluid communication with the second ends of at least two of the tubular reactors and a second end extending into the separation zone transports product and catalyst to the separation zone. A catalyst return in fluid communication with the separation zone and the catalyst retention zone returns catalyst to the catalyst retention zone.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Continuation of copending application Ser. No.10/125,468 filed Apr. 18, 2002, the contents of which are herebyincorporated by reference in its entirety.

BACKGROUND OF THE INVENTION

This invention relates generally to a process and apparatus forcontacting fluidized catalyst with hydrocarbon streams. Morespecifically, this invention relates to processes for upgradinghydrocarbon feeds in a discrete reactor vessel.

DESCRIPTION OF THE PRIOR ART

The FCC process is carried out by contacting the starting materialwhether it be vacuum gas oil, reduced crude, or another source ofrelatively high boiling hydrocarbons with a catalyst made up of finelydivided or particulate solid material. The catalyst is transported in afluid-like manner by passing gas or vapor through it at sufficientvelocity to produce a desired regime of fluid transport. Contact of theoil with the fluidized material catalyzes the cracking reaction. Thecracking reaction deposits coke on the catalyst. Catalyst exiting thereaction zone is spoken of as being “spent”, i.e., partially deactivatedby the deposition of coke upon the catalyst. Coke is comprised ofhydrogen and carbon and can include other materials in trace quantitiessuch as sulfur and metals that enter the process with the startingmaterial. Coke interferes with the catalytic activity of the spentcatalyst by blocking acid sites on the catalyst surface where thecracking reactions take place. Spent catalyst is traditionallytransferred to a stripper that removes adsorbed hydrocarbons and gasesfrom catalyst and then to a regenerator for purposes of removing thecoke by oxidation with an oxygen-containing gas. An inventory ofcatalyst having a reduced coke content, relative to the spent catalystin the stripper, hereinafter referred to as regenerated catalyst, iscollected for return to the reaction zone. Oxidizing the coke from thecatalyst surface releases a large amount of heat, a portion of whichescapes the regenerator with gaseous products of coke oxidationgenerally referred to as flue gas. The balance of the heat leaves theregenerator with the regenerated catalyst. The fluidized catalyst iscontinuously circulated between the reaction zone and the regenerationzone. The fluidized catalyst, as well as providing a catalytic function,acts as a vehicle for the transfer of heat from zone to zone. The FCCprocesses, as well as separation devices used therein are fullydescribed in U.S. Pat. No. 5,584,985 and U.S. Pat. No. 4,792,437, thecontents of which are hereby incorporated by reference. Specific detailsof the various contact zones, regeneration zones, and stripping zonesalong with arrangements for conveying the catalyst between the variouszones are well known to those skilled in the art.

The FCC reactor cracks gas oil or heavier feeds into a broad range ofproducts. Cracked vapors from the FCC unit enter a separation zone,typically in the form of a main column, that provides a gas stream, agasoline cut, light cycle oil (LCO) and clarified oil (CO) whichincludes heavy cycle oil (HCO) components. The gas stream may includedry gas, i.e., hydrogen and C₁ and C₂ hydrocarbons, and liquefiedpetroleum gas (“LPG”), i.e., C₃ and C₄ hydrocarbons. LPG and dry gas arenot as valuable as naphtha in many regions. However, LPG is morevaluable than naphtha in some regions. The gasoline cut may includelight, medium and heavy gasoline components. A major component of theheavy gasoline fraction comprises condensed single ring aromatics. Amajor component of LCO is condensed bicyclic ring aromatics.

Subjecting product fractions to additional reactions is useful forupgrading product quality. The recracking of heavy product fractionsfrom the initially cracked FCC product is one example. Typically, inrecracking, uncracked effluent from a first riser of an FCC reactor isrecontacted with catalyst at a second location to cleave largermolecules down into smaller molecules. For example, U.S. Pat. No.4,051,013 discloses cracking both gasoline-range feed and gas oil feedin the same riser at different elevations. WO 01/00750 A1 disclosesintroducing gasoline feed and FCC feed at different elevations in ariser reactor, separating the cracked product and recycling portionsthereof back to the same riser reactor.

U.S. Pat. No. 2,921,014, U.S. Pat. No. 3,161,582, U.S. Pat. No.5,176,815 and U.S. Pat. No. 5,310,477 all disclose cracking a primaryhydrocarbon feed in a riser of an FCC unit and cracking a secondaryhydrocarbon feed in a reactor into which the riser exits. As a result,both cracked products mix in the reactor, to some extent, which couldnegate the incremental upgrade resulting from cracking the secondaryhydrocarbon feed, particularly if it is a fraction of the crackedprimary hydrocarbon feed. U.S. Pat. No. 2,956,003 avoids thisdisadvantage to some extent by separating the vapor product and thecatalyst of the riser effluent and contacting the secondary hydrocarbonfeed in a reactor which receives the separated catalyst.

FCC units employing two risers are known. U.S. Pat. No. 5,198,590, U.S.Pat. No. 4,402,913, U.S. Pat. No. 4,310,489, U.S. Pat. No. 4,297,203,U.S. Pat. No. 3,799,864, U.S. Pat. No. 3,748,251, U.S. Pat. No.3,714,024 and US 2002/0003103 A1 disclose two riser FCC units in whichfeeds are predominantly cracked in both risers. In these patents, bothrisers communicate with the same recovery conduit and/or reactorpermitting commingling of gaseous products. In U.S. Pat. No. 5,730,859,all of the effluent from one riser is fed to the other riser, withoutfirst undergoing a product separation. U.S. Pat. No. 4,172,812 teachesrecracking all or a part of cracked product from a riser of an FCC unitover a catalyst having a composition that is different from the catalystcomposition in the riser. U.S. Pat. No. 5,401,387 discloses crackingfeed in a fixed, fluidized or moving bed of shape selective catalyst andfeeding all or part of the cracked effluent mixed with heavier feed toan FCC unit. In U.S. Pat. No. 5,944,982, although both risers terminatein the same reactor vessel, gaseous products from each riser areisolated from the other.

Two types of flow regimes have been used in secondary reactors.Transport flow regimes are typically used in FCC riser reactors. Intransport flow, the difference in the velocity of the gas and thecatalyst, called the slip velocity, is relatively low, typically lessthan 0.3 m/s (1.0 ft/s) with little catalyst back mixing or hold up.Slip velocity is calculated by the following formula: $\begin{matrix}{v_{s} = {\frac{u_{g}}{ɛ} - u_{s}}} & (1)\end{matrix}$where v_(s) is the slip velocity, u_(g) is the superficial gas velocity,u_(s) is the catalyst velocity and ε is the void fraction of thecatalyst. Another way to characterize flow regimes is by slip ratiowhich is the ratio of actual density in the flow zone to the non-slipdensity in the flow zone. The non-slip density is calculated by theratio of catalyst flux to the superficial gas velocity: $\begin{matrix}{\rho_{n\quad s} = \frac{\omega_{c}}{u_{g}}} & (2)\end{matrix}$where ρ_(ns) is the non-slip density in the flow zone, ω_(c) flux of thecatalyst and u_(g) is the superficial gas velocity. Catalyst flux is themass flow rate of catalyst per cross-sectional area of the reactor. Theslip ratio is proportional to the hold up of catalyst in the flow zone.Typically, a slip ratio for a transport flow regime does not reach 2.5.Consequently, the catalyst in the reaction zone maintains flow at a lowdensity and very dilute phase conditions. The superficial gas velocityin transport flow is typically greater than 3.7 m/s (12.0 ft/s), and thedensity of the catalyst is typically no more than 48 kg/m³ (3 lb/ft³)depending on the characteristics and flow rate of the catalyst andvapor. In transport mode, the catalyst-vapor mixture is homogeneouswithout vapor voids or bubbles forming in the catalyst phase.

Bubbling bed secondary reactors are also known. In a bubbling bed,fluidizing vapor forms bubbles that ascend through a discernible topsurface of a dense catalyst bed. Only catalyst entrained in the vaporexits the reactor with the vapor. The superficial velocity of the vaporis typically less than 0.5 m/s (1.5 ft/s) and the density of the densebed is typically greater than 640 kg/m³ (40 lb/ft³) depending on thecharacteristics of the catalyst. The mixture of catalyst and vapor isheterogeneous with pervasive vapor bypassing of catalyst.

Intermediate of bubbling beds and transport flow regimes are turbulentbeds and fast fluidized regimes. U.S. Pat. No. 4,547,616 discloses aturbulent flow regime for oxygenate conversion. In a turbulent bed, themixture of catalyst and vapor is not homogeneous. The turbulent bed hasa dense catalyst bed with elongated voids of vapor forming within thecatalyst phase, and the surface is less discernible. Only entrainedcatalyst leaves with the vapor and the catalyst density is not quiteproportional to its elevation within the reactor. U.S. Pat. No.6,166,282 discloses a fast fluidized flow regime used for oxygenateconversion. In a fast fluidized regime, there is no dense catalyst bed.Instead, the catalyst and vapor phases are homogeneous. Catalyst exitsthe reaction zone a small amount slower than the vapor exiting thereaction zone. Hence, for a fast fluidized flow regime the slip velocityis typically greater than or equal to 0.3 m/s (1.0 ft/s) and the slipratio is greater than or equal to 2.5 for most FCC catalysts. Fastfluidized beds have been used in FCC combustors for regeneratingcatalyst and in coal gasification.

U.S. Pat. No. 3,928,172 teaches an FCC unit with a secondary reactor.Gas oil is cracked in a riser of the FCC unit with unregenerated spentcatalyst under transport flow conditions. A heavy naphtha fraction ofthe cracked gas oil, boiling between 127° and 232° C. (260° and 450°F.), produced in the riser is recracked in the secondary reactor overregenerated catalyst in a bubbling bed. The spent catalyst used in theriser comes from the dense bubbling bed in the secondary reactor.

U.S. Pat. No. 5,346,613, U.S. Pat. No. 5,451,313, U.S. Pat. No.5,455,010, U.S. Pat. No. 5,597,537, U.S. Pat. No. 5,858,207, U.S. Pat.No. 6,010,618 and U.S. Pat. No. 6,113,776 all disclose recyclingunregenerated spent catalyst that has been contacted with feed back tothe reactor section to be mixed with regenerated catalyst andrecontacted with new feed for purposes of lowering the temperature ofthe catalyst to contact the feed. U.S. Pat. No. 5,965,012 disclosescontacting a first feed with mixed unregenerated spent catalyst andregenerated catalyst in a conduit that exits into a mixing vessel. Vaporproduct is removed from the mixing vessel and the mixed catalystcontacted with the first feed is further mixed with other unregeneratedspent catalyst and the further mixed catalyst contacts a second feed inan FCC riser.

In gasoline production, many governmental entities are restricting theconcentration of olefins allowed in the gasoline pool. Reducing olefinconcentration without also reducing value is difficult because higherolefin concentrations typically promote higher Research Octane Numbers(RON) and Motor Octane Numbers (MON), but the latter to a lesser extent.Octane value or Road Octane Number is the average of RON and MON. Merelysaturating olefins typically yields normal paraffins which typicallyhave low octane value. Additionally, saturation requires the addition ofhydrogen, which is expensive and in some regions, difficult to obtain.

US 2004/0140246 A1 discloses reacting naphtha over FCC catalyst in aseparate reactor of an FCC unit under conditions that promote hydrogentransfer reactions without having to add hydrogen. The hydrogen transferreactions promote reformulation of olefins to isoparaffins andaromatics. The reformulation to isoparaffins reduces octane value, butthe reduction is not as great because the resulting paraffins areisoparaffins, which have a greater octane value than normal paraffins.Moreover, the reformulation of larger olefins to aromatics operates toboost the octane value, thereby offsetting any loss due to saturation ofolefins to isoparaffins.

EP 1 046 696 A2 and EP 1 046 695 A2 disclose contacting feedstock withcatalyst in a first reaction zone of a riser and quenching the effluentof the first reaction zone with a medium which can include a regeneratedand cooled catalyst or naphtha. The quenched effluent from the firstreaction zone passes to a second reaction zone of the same riser whichmay have a greater diameter than the first reaction zone underconditions that promote isomerization and hydrogen transfer reactions.The feedstock to the first reaction zone may include naphtha. Anarticle, X. Youhao, Z. Jiushun & L. Jun, “A Modified FCC Process MIP forMaximizing Iso-Paraffins in Cracked Naphtha”, PETROLEUM PROCESSING &PETROCHEMICALS (August 2000), discloses reacting gasoline over an FCCspent catalyst. These disclosures report high reductions in olefinconcentration and increases in isoparaffin and aromatics concentration.

Feedstocks for FCC units typically include organic sulfur and nitrogen.During FCC operation, some of the organic sulfur and nitrogen areconverted to hydrogen sulfide and ammonia, which are easily removed.Some, however, are converted to coke, lighter sulfur and nitrogencompounds and mercaptans. This coke is then oxidized in the catalystregenerator to form sulfur oxides and nitrogen oxides. Stricterenvironmental limits on sulfur and nitrogen compound emissions haveprecipitated lower sulfur specifications for fuel products therebyraising interest in removing nitrogen and sulfur compounds from FCCgasoline. As demand for cleaner fuels and use of high sulfur and highnitrogen feedstocks increase, the need for sulfur and nitrogen removalfrom FCC gasoline will become even greater.

U.S. Pat. No. 5,482,617 discloses a process for desulfurizing ahydrocarbon stream such as FCC naphtha containing sulfur by contactingit with an acidic catalyst in a bed. As much as 50 wt-% of the sulfur isconverted to hydrogen sulfide.

WO 01/00751 A1 discloses a conversion process for reducing olefins,sulfur and nitrogen concentrations in gasoline. Preheated gasoline iscontacted with catalyst having no more than 2.0 wt-% carbon depositionand a temperature of below 600° C. The gasoline product has an olefincontent reduced to below 20 wt-% and sulfur and nitrogen contents arealso reduced. It appears that the regenerated catalyst is cooled in acatalyst cooler to achieve the temperature below 600° C. beforecontacting the gasoline feed.

Cooling catalyst with a cooler can be inefficient. The crackingreactions in an FCC reactor are endothermic. In addition, heat isrequired to vaporize feed for the secondary reactor. Hence, cooling thecatalyst withdraws heat from the process that must be replaced in orderto vaporize the feed to the separate reactor. The heat is replaced bygenerating more coke on catalyst to fuel the regenerator with thesacrifice of valuable product.

It is an object of the present invention to provide a method forenhancing the quality of product obtain in a hydrocarbon conversionprocess by reacting it under conditions that assure thorough mixing ofcatalyst and feed.

SUMMARY OF THE INVENTION

It has now been discovered that contacting hydrocarbon feed withcatalyst in a reactor under fast fluidized flow conditions is ideal forpromoting hydrogen transfer reactions and controlled catalytic crackingreactions. Flow conditions may descend into a turbulent bed regime butwith less effectiveness compared to fast fluidized flow. The secondaryreactor may be incorporated into an FCC reactor and utilize catalysttherefrom. The homogeneous mixing of catalyst and feed vapor phasesprovides for sufficient catalyst to feed contact while avoiding unduegeneration of coke and dry gas.

Accordingly, in one embodiment, the present invention relates to aprocess for cracking and further treating hydrocarbons. The processcomprises contacting a first hydrocarbon feed stream with catalyst toyield cracked hydrocarbons and spent catalyst. The spent catalyst isthen separated from the cracked hydrocarbons in a separator section. Atleast a portion of the spent catalyst is regenerated to provideregenerated catalyst. Then a second hydrocarbon feed stream is contactedwith regenerated catalyst in a reaction zone under flow conditionsincluding a superficial vapor velocity of greater than or equal to 0.6m/s (1.8 ft/s) and a slip ratio of greater than or equal to 2.5 to yieldupgraded hydrocarbons and spent catalyst.

In another embodiment, the present invention relates to a process fortreating a hydrocarbon stream including at least a portion of effluentfrom an FCC reactor. The process comprises contacting the hydrocarbonstream with catalyst in a reaction zone under conditions that promotehomogeneous mixing of catalyst and hydrocarbons and a slip ratio ofgreater than or equal to 3.0. Such contacting yields an upgradedhydrocarbon stream and spent catalyst.

In a further embodiment, the present invention relates to an apparatusfor the contacting of hydrocarbons with catalyst. The apparatuscomprises a reactor vessel including at least one reactor and a diluentnozzle communicating with the reactor vessel. A feed nozzle communicateswith the reactor at a first end of the reactor and a second end of thereactor has a reduced cross-sectional area relative to the reactor.Lastly, a separator vessel including a transport conduit communicateswith the second end of the reactor and the transport conduit has adischarge opening communicating with the separator vessel.

Additional objects, embodiment and details of this invention can beobtained from the following detailed description of the invention.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a sectional, elevational, schematical view of an FCC reactorincorporating a secondary reactor in accordance with the presentinvention.

FIG. 2 is a sectional, plan view taken along segment 2-2 in FIG. 1.

DETAILED DESCRIPTION OF THE INVENTION

We have found that when reacting hydrocarbons produced by an FCC processin a secondary reactor they can potentially crack down to dry gas andLPG which are less valuable than naphtha in many regions. Reducedcatalyst temperature will suppress thermal cracking of the naphthaolefins to dry gas and LPG which allows for higher overall naphthayields. In one embodiment, the present invention recycles spent catalystto the secondary reactor to be recontacted with feed. The recycled spentcatalyst mixes with regenerated catalyst from a regenerator to lower theoverall mixed catalyst temperature.

Mixing the recycled spent catalyst with regenerated catalyst is apreferred method of cooling catalyst compared to use of an indirect heatexchange catalyst cooler, which operates to withdraw heat from thesystem that must be replaced by the generation and burning of more coke.Even if heat is conserved by heating feed to a reactor by indirect heatexchange with hot spent catalyst in the catalyst cooler, this means ofheating would still be less efficient and prone to cause coking.

The use of recycled, coke-bearing, spent catalyst also benefitsoperation by tempering catalyst activity. The strong acid sites of theFCC catalyst are suppressed through formation of coke. Consequently, thegasoline reformulation activity relative to the cracking activity isincreased. Moderating the catalyst activity permits higher naphthayields to be reached at the expense of production of dry gas and LPG.

We have also discovered that better mixing between the feed and thecatalyst promotes hydrogen transfer reactions and catalytic crackingreactions while reducing the undesirable generation of coke and dry gas.While not wishing to be bound by any particular theory, it is believedthat better mixing allows the hydrocarbon to reach the catalyst surfacemore rapidly which enhances hydrogen transfer and catalytic crackingreactions. This allows the total contact time and/or catalyst to feedratio to be reduced which minimizes dry gas and coke yield at a givenconversion. Hence, reactor conditions that are more vigorous than abubbling bed such as a fast fluidized flow regime will provide betterproduct yields.

The present invention may be described with reference to threecomponents: an FCC reactor 10, a regenerator 50 and a secondary reactor70. Although many configurations of the present invention are possible,one specific embodiment is presented herein by way of example. All otherpossible embodiments for carrying out the present invention areconsidered within the scope of the present invention.

In the embodiment of the present invention in FIG. 1, the FCC reactor 10comprises a reactor vessel 12 in the form of a conduit that ispreferably vertical which is also known as a riser. The reactor vessel12 extends upwardly through a lower portion of a separator vessel 14 asin a typical FCC arrangement. The reactor vessel 12 preferably has avertical orientation within the separator vessel 14 and may extendupwardly through the bottom of the separator vessel 14 or downwardlyfrom the top of the separator vessel 14. The reactor vessel 12terminates in a disengagement section 16 of the separator vessel 14 atswirl tubes 18. A hydrocarbon feed stream is fed to the riser at anozzle 20 which is contacted and vaporized by hot regenerated catalystfluidized by a gas such as steam from a nozzle 22. The catalyst cracksthe hydrocarbon feed stream and a mixture of spent catalyst particlesand gaseous cracked hydrocarbons exit discharge openings in the swirltubes 18 into the disengagement section 16. Tangential discharge ofgases and spent catalyst from the swirl tubes 18 produces a swirlinghelical motion about the interior of the disengagement section 16,causing heavier catalyst particles to fall into a dense catalyst bed 24and a mixture of gaseous cracked hydrocarbons and entrained spentcatalyst particles to travel up a gas recovery conduit 26 and enter intocyclones 28. In the cyclones 28, centripetal force imparted to themixture induces the heavier entrained catalyst particles to fall throughdiplegs 30 of the cyclone 28 into a dense catalyst bed 32 at the bottomof the separator vessel 14. The gases in the cyclones 28 more easilychange direction and begin an upward spiral with the gases ultimatelyexiting the cyclones 28 through outlet pipes 34. Cracked gases leave theseparator vessel 14 though an outlet conduit 36. The cracked gases areoptionally sent via a line 46 to a further separation (not shown) toremove any light loading of catalyst particles and then tofractionation. Spent catalyst particles in the dense catalyst bed 32enter the disengagement section 16 through windows 38 where they joinspent catalyst particles in the dense catalyst bed 24 in a strippingsection 40 of the disengagement section 16. The spent catalyst particlesare stripped of entrained cracked vapors over baffles 42 with astripping medium such as steam entering from at least one nozzle 44. Thestripped cracked vapors travel up to the gas recovery conduit 26 wherethey are processed with other cracked product vapors.

Stripped spent catalyst from the stripping section 40 of the separatorvessel 14 of the FCC reactor 10 travels through a spent catalyst pipe 48regulated by a control valve 49 and preferably into the regenerator 50.In an embodiment, some of the stripped catalyst may be delivered to thesecondary reactor 70. In the regenerator 50, stripped spent catalyst issubjected to hot oxygen-containing gas such as air from a distributor53. Coke is burned from the spent catalyst as the catalyst is heated.Regenerated catalyst collects in a dense catalyst bed 58 whereasentrained catalyst is removed from regenerator effluent gases incyclones 60 and 62. Flue gas exits the cyclone 62 through an outlet pipe64 to exit the regenerator 50 through an outlet 66. Regenerated catalystfrom the dense catalyst bed 58 travels through a first regeneratedcatalyst pipe 68 regulated by a control valve 69 into the reactor vessel12 where it is fluidized and contacted with fresh feed. Regeneratedcatalyst also exits the regenerator 50 through a second regeneratedcatalyst pipe 52 regulated by a control valve 54 into a mixing pot 72 ofand at the base of the secondary reactor 70.

The contacting of feed and catalyst occurs in a reactor vessel 74 of thesecondary reactor 70. Control valves 54,80 govern the rate of catalystcirculation to the reactor vessel 74. If the reactor vessel 74 includesa catalyst bed, the catalyst circulation rate influences the height ofthe catalyst bed in the reactor vessel 74. The height of the catalystbed in the reactor vessel 74 of the secondary reactor 70 influences theweight hourly space velocity (WHSV) of reactants through the reactorvessel 74. For example, if a greater WHSV is desired, the control valves54, 80 would be closed relatively more to reduce the height of catalystin a dense catalyst bed and the ratio of catalyst to oil. On the otherhand, if smaller WHSV is desired, the control valves 54, 80 would beopened relatively more to increase the level of catalyst in the densecatalyst bed 82 and the ratio of catalyst to feed. Relative settings ofthe control valves 54, 80 are independently adjusted also to obtain thedesired temperature and mixture of the catalyst in the dense catalystbed 82 that will contact the feed in reactors 84.

The cracked product stream in the line 46 from the FCC reactor 10,relatively free of catalyst particles and including the stripping fluid,is transferred to a fractionator main column which is not shown in thedrawings. One or more cuts from the main column is preferably sent tothe secondary reactor 70 to be contacted with the catalyst therein. Inone embodiment, a cut from the main column such as a light cycle oil cut(LCO) may be hydrotreated in a hydrotreating reactor before it is sentto the secondary reactor 70 for cracking.

In a preferred embodiment, the secondary reactor 70 includes the reactorvessel 74 and a separator vessel 76. Spent mixed catalyst is deliveredby a recycle spent mixed catalyst pipe 78 governed by the control valve80 and regenerated catalyst is delivered by the second regeneratedcatalyst pipe 52 governed by the control valve 54 to the mixing pot 72.Fluidizing medium such as steam delivered by a line 77 to a diluentnozzle 79 to fluidize the mixed catalyst in the mixing pot 72 andgenerates the dense catalyst bed 82 having an upper level 83. The mixingpot 72 enables adequate mixing and temperature equilibration of spentand regenerated catalyst before it is introduced to the feed. Theelevation of the upper level 83 will be proportional to the settings ofthe control valves 54, 80 relative to fully open. Although one reactor84 may be used for carrying out the purposes of this invention, theprovision of a plurality of reactors 84 with dedicated nozzles 88 andcontrol valves 89 to govern the flow rate of feed to each of thereactors 84 may offer flexibility over use of a single reactor 84 byincreasing the operating range of space velocity.

FIG. 2 is an upwardly looking cross-sectional view taken along segment2-2 of the reactor vessel 74. FIG. 2 shows a preferred embodimentincluding four reactors 84 in the reactor vessel 74.

Hydrocarbon feed which may be a fraction or all of the effluent in theline 46 from the FCC reactor 10 is distributed by a line 86 through thenozzles 88 to the respective reactors 84 in the reactor vessel 74. Thereactors 84 may be tubular with an open bottom end communicating withthe dense catalyst bed 82 in the reactor vessel 74. Feed entering thereactors 84 preferably pulls mixed catalyst from the dense catalyst bed82 into the reactors where contacting occurs. The amount of mixedcatalyst pulled into the reactor 84 for a given flow rate of feed willbe proportional to the height of upper level 83 of the dense catalystbed 82 which is controlled by control valves 54, 80. Hence, the ratio ofcatalyst to feed, the space velocity and the density in the reactor canall be controlled by controlling the elevation of the upper level 83 ofthe dense catalyst bed 82 and/or controlling or eliminating the flowrate of the feed through one or more of the nozzles 88.

A top end 85 of the reactors 84 has a reduced cross-sectional area thatmay take the form of a frusto-conical section. The reducedcross-sectional area of the top end 85 serves to accelerate the mixtureof vapor product and spent mixed catalyst as they exit the reactorvessel 74 and enter outlet conduits 90. The outlet conduits 90communicate the reactor vessel 74 with a transport conduit 92 in theform of a riser. The outlet conduits 90 may be supplanted by a directconnection of the reactor 84 to the transport conduit 92. The outletconduits 90 should all have a smaller cross-sectional area than therespective reactor 84 and the transport conduit 92 preferably has across-sectional area that is less than the aggregate cross-sectionalarea of all of the reactors 84 that feed the transport conduit 92.Consequently, upon leaving the reactor 84, the product vapor and spentmixed catalyst accelerate into a transport mode, thus giving the spentmixed catalyst and product vapor little time to further react or crackinto undesirable products. Entering transport mode also preventscatalyst from falling out of entrainment with the product vapor.

Spent mixed catalyst and product ascend from the reactor vessel 74through the transport conduit 92 to the separator vessel 76. The spentmixed catalyst and vapor product exit through discharge openings 94 inswirl tubes 96 to effect a primary, centripetal separation of spentmixed catalyst from the vapor product. Separated spent mixed catalystsettles into a dense bed 98 in the separator vessel 76. The spent mixedcatalyst in the separator vessel 76 is then preferably stripped over aseries of baffles 100 by use of a stripping medium such as steamentering through stripping nozzles 102 in a stripping section 103 of theseparator vessel 76. A first portion of the stripped spent mixedcatalyst exits the separator vessel 76 through a spent mixed catalystpipe 104 at a flow rate governed by a control valve 106. The strippedspent mixed catalyst enters a riser 108 which distributes it to the bed58 in the regenerator 50. Alternatively, the spent mixed catalyst pipe104 may deliver spent mixed catalyst to the FCC reactor 10. A secondportion of the stripped spent mixed catalyst is withdrawn through therecycle spent mixed catalyst pipe 78 at a flow rate governed by thecontrol valve 80 and is delivered to the mixing pot 72 where it is mixedwith regenerated catalyst delivered from the second regenerated catalystpipe 52. Product vapors and entrained catalyst are withdrawn from theseparator vessel 76 through an outlet pipe 110 and are delivered to anexternal cyclone separator 112. Alternatively, the cyclone separator 112could be installed inside the separator vessel 76. The entrainedcatalyst is centripetally separated from product vapors in the cycloneseparator 112. Separated catalyst exits through a dipleg 116 and isreturned to the reactor vessel 74 at a flow rate governed by a controlvalve 118. Product vapors in a line 120 are withdrawn from the cycloneseparator 112 through an outlet 114 and sent to further productprocessing.

It is important to not mix the vaporous product from the secondseparator vessel 14 with the vaporous product from the first separatorvessel 76. Such mixing could negate upgrading of the product wrought inthe secondary reactor 70. This is especially true when the secondaryreactor 70 processes feed that is derived from the FCC reactor 10.Hence, the produce in the line 120 is preferably isolated from theproduct in the line 46. Additionally, it is preferable to keep pipes 48,52, 68, 104 separate.

This invention can employ a wide range of commonly used FCC catalysts.These catalyst compositions include high activity crystalline aluminasilicate or zeolite containing catalysts. Zeolite catalysts arepreferred in FCC operation because of their higher intrinsic activityand their higher resistance to the deactivating effects of hightemperature exposure to steam and exposure to the metals contained inmost feedstocks. Zeolites are usually dispersed in a porous inorganiccarrier material such as silica, aluminum, or zirconium. These catalystcompositions may have a zeolite content of 30% or more. Zeolitesincluding high silica-to-alumina compositions such as LZ-210, ST-5 andZSM-5 type materials are preferred when lighter products are desiredfrom either the FCC or secondary reactor. Another particularly usefultype of FCC catalysts comprises silicon substituted aluminas. Asdisclosed in U.S. Pat. No. 5,080,778, the zeolite or silicon enhancedalumina catalysts compositions may include intercalated clays, alsogenerally known as pillared clays. The preferred catalysts for thepresent invention include USY zeolites.

Typically, the catalyst circulation rate through the reactor vessel 12of the FCC reactor 10 and the input of feed and any lift gas that entersthe reactor vessel 12 will be operated to produce transport conditionswith a flowing density of less than 48 kg/m³ (3 lbs/ft³) and an averagesuperficial velocity of more than 3.7 up to 31 m/s (12 to 93 ft/s). Inthe FCC reactor 10, catalyst will usually contact the hydrocarbons in acatalyst to oil ratio in a range of from 3 to 8, and more preferably ina range of from 4 to 6. The length of the reactor vessel 12 will usuallybe set to provide a residence time of between 0.5 to 10 seconds at theseaverage flow velocity conditions. Other reaction conditions in thereactor vessel 12 usually include a temperature of from 468° to 566° C.(8750 to 1050° F.).

Feeds suitable for processing in the FCC reactor 10 include conventionalFCC feedstocks or higher boiling hydrocarbon feeds. The most common ofthe conventional feedstocks is a vacuum gas oil which is typically ahydrocarbon material having a boiling range of from 343° to 552° C.(650° to 1025° F.) and is prepared by vacuum fractionation ofatmospheric residue. Such fractions are generally low in coke precursorsand heavy metals which can deactivate the catalyst.

The effluent from the FCC reactor 10 in the line 46 may be processed ina main column (not shown) into fractions that include a light off-gasstream, a light gasoline liquid stream, a heavy gasoline stream, a lightcycle oil (“LCO”) stream, a heavy cycle oil (“HCO”) stream and aclarified oil (“CO”) stream. The light gasoline or light naphthafraction preferably has an initial boiling point (IBP) at or above about−5° C. (23° F.) in the C₄ range and an end point (EP) at a temperaturegreater than or equal to 127° C. (260° F.). The boiling points for thesefractions are determined using the procedure known as ASTM D86-82. Theheavy gasoline or heavy naphtha fraction has an IBP at or above 127° C.(260° F.) and an EP preferably between 204° and 232° C. (4000 and 450°F.), particularly at 216° C. (420° F.). The LCO stream has an IBP atabout the EP temperature of the heavy gasoline and an EP in a range of2600 to 371° C. (500° to 700° F.) and preferably 288° C. (550° F.). TheHCO stream has an IBP of the EP temperature of the LCO stream and an EPin a range of 3710 to 427° C. (700° to 800° F.), and preferably about399° C. (750° F.). The CO stream has an IBP of the EP temperature of theHCO stream and includes everything boiling at a higher temperature. Anyor all of these fractions may be treated in the secondary reactor 70.The hydrocarbon feed preferably comprises at least 50 wt-% naphtha.

Alternatively, the feed to the secondary reactor 70 may come from asource other than the effluent from the FCC reactor 10. Coker naphtha isone example of feed that could be independently derived and fed to thesecondary reactor 70.

In the secondary reactor 70, the predominant reaction may be cracking inwhich a hydrocarbon molecule is broken into two smaller hydrocarbonmolecules, so that the number of carbon atoms in each moleculediminishes. Alternatively, the predominant reaction in the secondaryreactor 70 may be a hydrogen-transfer reaction such as reformulation orisomerization in which the structures of the molecules are changed butthe number of carbon atoms in each molecule does not change.

Olefins, naphthenes and cyclo-olefins are reformulated into paraffins,aromatics and some naphthenes as shown in formulas (3), (4), (5) and(6).

Olefins have a higher octane value than their paraffinic counterpart.Hence, the conversion of olefins to paraffins typically degrades octanevalue. When the olefins cyclitize to become aromatics as shown informulas (3) and (4) and when cyclo-olefins aromaticize to yieldaromatics as in formula (5), they donate much hydrogen. Other olefinspick up the hydrogen to become paraffins as shown in formula (6). In thepresent invention using the secondary reactor 70, normal olefins andisoolefins predominantly reformulate to isoparaffins which carry ahigher octane rating than normal paraffins. Additionally, aromatics alsoboost the octane rating of the product. Because the isoparaffins andaromatics have a high octane rating, the hydrogen transfer reformulationin the secondary reactor 70 maintains the high octane ratings despitethe typical octane rating decline that accompanies conversion of olefinsto paraffins. Accordingly, the hydrogen-transfer reactions in thesecondary reactor 70 which yield more isoparaffins and aromatics aresuperior to a process that saturates the olefins into normal paraffins.Advantageously, the hydrogen transfer reactions are performed withoutthe addition of hydrogen, which can be expensive and difficult toobtain.

The reaction in the secondary reactor 70 is preferably conducted withthe same catalyst circulated through the regenerator 50 and the FCCreactor 10. Of course, if the secondary reactor 70 stands alone withoutincorporation into an FCC reactor, the catalyst in the secondary reactor70 need not be circulated through an FCC reactor.

When hydrogen-transfer reactions are desired to predominate overcracking reactions in the secondary reactor 70, high rare earth contentY zeolites are preferred. The term “high rare earth content” denotesgreater than about 2.0 wt-% rare earth oxide on the zeolite portion ofthe catalyst. High rare earth content Y zeolites such as USY zeolite mayhave as much as 4 wt-% rare earth. The high rare earth content promoteshydrogen transfer by increasing adjacent acid site density on thecatalyst. Strongly acidic catalyst sites on the catalyst promotecracking. Y zeolites with low rare earth content can still effectivelypromote hydrogen transfer but with longer reactor residence times. Whencracking reactions are desired to predominate over hydrogen transferreactions in the secondary reactor 70, low rare earth Y zeolitecatalysts are preferred which have a rare earth oxide content of 2.0wt-% or less. Additives, such as sulfur-reducing additives, may be addedto the catalyst. It is anticipated that such additives may experienceenhanced effectiveness in the secondary reactor 70 for longer residencetimes.

In an embodiment of the present invention, the secondary reactor 70 maybe operated at bubbling bed conditions. In such an embodiment, separatereactors 84 may not be necessary. In an additional embodiment, thesecondary reactor 70 may be operated at transport conditions. In thisembodiment, the reactor vessel 74 or the reactors 84 would likelycomprise a conduit such as a riser. However, it is preferred to contactcatalyst and feed in the secondary reactor 70 under fast fluidized flowconditions. Even under fast fluidized conditions, the mixed catalystpreferably accumulates in the dense catalyst bed 82 operated at bubblingbed conditions until it is pulled into one of the reactors 84. Hence,the fluidizing medium is delivered by the nozzle 79 preferably togenerate a superficial velocity in the reactor vessel 74 of less than0.5 m/s (1.5 ft/s) and a density of greater than 480 kg/m³ (30 lb/ft³).Once the hydrocarbon feed and the catalyst enter the reactors 84, theypreferably flow in a fast fluidized flow regime. Hence, the mixture ofcatalyst and vaporized feed is homogeneous leaving no discerniblesurface of a bed in the reactor 84. Thoroughly mixing catalyst and vaporphases reduces the generation of coke. To achieve the fast fluidizedregime, the superficial velocity of the vaporized feed from the nozzles88 should preferably be between 1.3 and 3.7 m/s (4 and 12 ft/s) and thedensity in the reactor 84 preferably should be between 48 and 320 kg/m³(3 and 20 lb/ft³) for typical FCC catalysts. Adequate, but less thoroughmixing, may occur even in a turbulent bed. However, the superficialvapor velocity should always be at least 0.6 m/s (1.8 ft/s) and thedensity in the reactor should never exceed 480 kg/m³ (30 lb/ft³) toavoid approaching a less desirable bubbling bed condition in thereactors 84 for typical FCC catalysts. In a fast fluidized flow regime,the slip ratio is between 2.5 and 10 and preferably at or above 3.0. Theslip velocity is high but does not extend into the transport mode ofless than 0.3 m/s (1.0 ft/s) and is preferably kept above or at 0.5 m/s(1.5 ft/s) for typical FCC catalysts to provide for back mixing to occurand adequate time for hydrogen transfer reactions to occur. The slipvelocity of greater than or equal to 0.3 m/s (1.0 ft/s) will not be ameaningful parameter for a fast fluidized flow regime at highsuperficial gas velocities such as those well into the typical transportrange. To assure that the effluent from the reactors 84 enters transportmode to terminate the reactions, the outlet conduits 90 and/or thetransport conduit 92 should be dimensioned to make attainable asuperficial velocity of greater than 3.7 m/s (12 ft/s), a flowingdensity of less than 48 kg/m³ (3 lb/ft³) and a slip ratio of less than2.5.

The foregoing variable ranges for bed and flow regimes are based on thecharacteristics of typical FCC catalysts and naphtha range vaporizedfeed. These ranges may vary if different catalysts or feeds withdifferent molecular weights are used in the present invention.

If hydrogen-transfer reactions are intended to predominate over crackingreactions in the secondary reactor 70, the WHSV should typically rangefrom 1 to 25 hr⁻¹ and the temperature should range from 399° to 510° C.(750° to 950° F.). Cracking reactions will be more frequent if both thespace velocity and temperature are at the high ends of these ranges.However, hydrogen transfer reactions may still predominate at higherspace velocity offset by lower temperature in these ranges and viceversa. If cracking reactions are to predominate over hydrogen-transferreactions in the secondary reactor 70, the WHSV should typically rangefrom 15 to 50 hr⁻¹ and the temperature should range from 482° to 649° C.(900° to 1200° F.). Hydrogen transfer reactions will be more frequent ifboth the space velocity and temperature are at the low ends of theseranges. However, cracking reactions may still predominate at lowertemperature offset by higher space velocity in these ranges and viceversa. To ensure controlled catalytic cracking reactions occur insteadof thermal cracking reactions, it is advantageous to keep thetemperature of the secondary reactor 70 below 566° C. (1050° F.). Hence,temperature control is important for both types of reactions.

Mixing regenerated catalyst and spent catalyst can reduce thetemperature of the mixed catalyst contacting the feed to the secondaryreactor 70 by 27° to 166° C. (50° to 300° F.) depending on theregenerator temperature and the ratio of recycled spent mixed catalystto regenerated catalyst mixed in the secondary reactor 70. It will bepreferred to operate the ratio of recycled spent catalyst to regeneratedcatalyst at 1:1 if predominant hydrogen transfer reactions are desired.

Spent catalyst from the separator vessel 14 of the FCC reactor 10 may berecycled to the secondary reactor 70 to be mixed with regeneratedcatalyst instead of or in addition to spent mixed catalyst from theseparator vessel 76. However, at least some regenerated catalyst willhave to be mixed with the spent catalyst or spent mixed catalyst to givethe mixed catalyst enough heat to vaporize the feed. Even using anexternal feed heater to vaporize the feed would be less efficient andprone to cause coking. In an alternative embodiment, spent mixedcatalyst from the separator vessel 76 may be transported to the FCCreactor 10 instead of to the regenerator 50.

If cracking reactions are desired to predominate over hydrogen transferreactions in the secondary reactor 70, less spent catalyst should berecycled to the secondary reactor 70. Hence, the ratio of regenerated tospent catalyst in the secondary reactor 70 will be higher. The resultinghigher reactor temperature and greater availability of uncoked acidsites on the regenerated catalyst will promote cracking reactions.

We have also found that higher pressures favor hydrogen transferreactions but also favor coke production. Hence, it is suitable tooperate the reactor at a pressure of between 69 and 207 kPa (10 and 30psig) and preferably between 83 and 138 kPa (12 and 20 psig).Additionally, to promote hydrogen transfer reactions, a catalyst to feedratio should be set between 2 and 8 and preferably between 4 and 7.Higher pressures also favor catalytic cracking but without significantcoke generation. Hence, when catalytic cracking reactions are desired topredominate in the secondary reactor 70, a suitable reactor pressure isbetween 138 and 276 kPa (20 and 40 psig). Moreover, because highertemperature is necessary to promote the catalytic cracking reactions,more catalyst will be circulated to the secondary reactor 70 to providesufficient heating. Hence, a catalyst to feed ratio between 5 and 11 andpreferably between 7 and 10 will be appropriate to promote catalyticcracking reactions.

The reformulation of the fraction from the main column by hydrogentransfer in the secondary reactor 70 reduces the concentrations oforganic sulfur and nitrogen compounds in the products. The reaction ofthe gasoline fraction in the secondary reactor 70 can lower sulfurconcentration in the reactor products by as much as 80 wt-% and nitrogenconcentration in the products by as much as 98 wt-%. Hence, the productsfrom the secondary reactor 70 will contain low concentrations of sulfurand nitrogen compounds. Leftover sulfur and nitrogen compounds can beremoved from the product by hydrotreating and taken off in the overheadof a finishing distillation column or by another suitable method ifnecessary to meet specifications.

1. An apparatus for the contacting of hydrocarbons with catalyst, said apparatus comprising: a plurality of tubular reactors, each of the tubular reactors having a first end into which a catalyst can be fed and a second end through which the catalyst can exit the tubular reactor; a catalyst retention zone provided to contain catalyst and from which catalyst can be fed to the tubular reactors; a separation zone provided to separate the catalyst from products of a reaction conducted in the apparatus; a transport conduit having a first end in fluid communication with the second ends of at least two of the tubular reactors and a second end extending into the separation zone; and a catalyst return in fluid communication with the separation zone and the catalyst retention zone.
 2. The apparatus of claim 1 wherein the plurality of tubular reactors is in fluid communication with a plurality of transport conduits.
 3. The apparatus of claim 1 wherein the plurality of tubular reactors is in fluid communication with a single transport conduit.
 4. The apparatus of claim 1 wherein the transport conduit further comprises a collection device for connecting the second ends of at least two of the riser reactors with the first end of the conduit.
 5. The apparatus of claim 1 wherein the smallest cross sectional area of the conduit is at least equal to the sum of the smallest cross sectional areas of the riser reactors in fluid communication therewith.
 6. The apparatus of claim 1 wherein the hydrocarbon conversion apparatus includes at least four riser reactors.
 7. The apparatus of claim 1 wherein riser reactors are contained within a common shell having a wall and first and second ends.
 8. The apparatus of claim 7 wherein the riser reactors and the wall of the shell define the catalyst retention zone.
 9. The apparatus of claim 8 wherein the shell defines the separation zone.
 10. The apparatus of claim 1 further comprising a fluid distributor in fluid communication with the catalyst retention zone, the fluid distributor being provided to feed a fluidizing agent to the catalyst retention zone to fluidize catalyst contained in the catalyst retention zone.
 11. The apparatus of claim 1 wherein the catalyst return is positioned externally to the tubular reactors.
 12. The apparatus of claim 1 further comprising a feed distributor including at least one feed head positioned adjacent to the first ends of the riser reactors wherein the feed distributor includes a flow control device which provides the feed to the riser reactors through the feed heads.
 13. The apparatus of claim 1 further comprising an impingement device positioned in the separation zone, the impingement device being provided to move catalyst away from the transport conduit to the catalyst return.
 14. The apparatus of claim 1 further comprising a catalyst regenerator in fluid communication with the apparatus.
 15. The apparatus of claim 13 further comprising a catalyst stripper in fluid communication with the apparatus and the catalyst regenerator.
 16. An apparatus for the contacting of hydrocarbons with catalyst, said apparatus comprising: a plurality of tubular reactors, each of the tubular reactors having a first end into which a catalyst can be fed and a second end through which the catalyst can exit the tubular reactors; a separation zone provided to separate the catalyst from products of a reaction conducted in the hydrocarbon conversion apparatus; a transport conduit extending into the separation zone and providing fluid communication between the second ends of at least two of the tubular reactors and the separation zone; and at least one catalyst return in fluid communication with the separation zone and the first ends of the tubular reactors, the catalyst return being provided to transfer the catalyst from the separation zone to the first ends of the tubular reactors.
 17. A process for the contacting of hydrocarbons with catalyst, said process comprising: (a) contacting a fluidizable catalyst with a fluidizing agent to fluidize the fluidizable catalyst; (b) feeding the catalyst and a feed to a plurality of tubular reactors, the plurality of tubular reactors being part of a single hydrocarbon conversion apparatus; (c) contacting the feed with the catalyst in the plurality of tubular reactors under conditions effective to convert the feed to a product; (d) consolidating the output of at least two of the tubular reactors in a conduit for transport into a separation zone; and (e) separating the catalyst from the product in the separation zone, the separation zone being in fluid communication with at least two of the tubular reactors through the conduit; (f) returning the catalyst from the separation zone to the plurality of tubular reactors; and (g) repeating steps (a) to (f).
 18. The process of claim 17 wherein the catalyst is returned to the plurality of tubular reactors through at least one catalyst return which is in fluid communication with the separation zone and the plurality of riser reactors.
 19. The process of claim 17 further comprising: regenerating at least a portion of the catalyst in a catalyst regenerator after separating the catalyst from the products to produce a regenerated catalyst; and returning the regenerated catalyst to at least one of the separation zone, the catalyst return, and the catalyst retention zone.
 20. The process of claim 17 further comprising: stripping at least a portion of the catalyst prior to regenerating the at least a portion of the catalyst. 